Distillate production from oxygenates in moving bed reactors

ABSTRACT

Systems and methods are provided for conversion of oxygenate feeds to distillate boiling range products using multiple moving bed reactor stages. The systems and methods allow for multiple stages to be used while avoiding the need for distillation or other boiling point based separation as the mixture of feed and effluent is passed between stages. Instead, a stripping gas is used to disengage the feed and effluent from the catalyst solids. In combination with an improved moving bed reactor design, this can allow substantially all of the feed and effluent from a first moving bed reactor stage to be passed into a second moving bed reactor stage, even when the feed and effluent include both vapor and liquid phase portions.

CROSS REFERENCE TO RELATED APPLICATIONS

This application claims priority to U.S. Provisional Application Ser.No. 62/865,615 filed Jun. 24, 2019, which is herein incorporated byreference in its entirety.

FIELD

Systems and methods are provided for formation of distillate boilingrange fuels, chemicals, and/or other products by upgrading of oxygenatesin moving bed reactors.

BACKGROUND

Moving bed reactors are a type of reactor that is potentially suitablefor reactions where a fluid phase is exposed to catalyst and/or othersolid particles at specified temperature and pressure conditions. Movingbed reactors provide an advantage due to the movement of the solidparticles. Because the solid particles flow within the reactor, it isrelatively easy to withdraw catalyst from a moving bed reactor on aperiodic basis to regenerate the catalyst.

Although moving bed reactors can facilitate catalyst regeneration,transfer of multiple phases between moving bed reactors can presentdifficulties. In particular, moving bed reactors are not conventionallyused in situations where a three phase, e.g., gas/liquid/solid,co-current flow is transferred from a first moving bed reactor to asecond moving bed reactor. Because each phase of the three-phase flowhas different flow properties, attempting to transfer a three-phase flowby conventional methods can result in uneven distribution of one or moreflow phases. Such uneven distribution can lead to substantially reducedactivity, temperature spikes, increased catalyst deactivation, and/orvarious other poor performance characteristics. Additionally,conventional methods of transferring a three-phase flow can suffer fromlimits on the ability to independently control the input flow rate ofeach phase into the reactor.

One option for overcoming the difficulties with managing co-current flowin a moving bed reactor is to use a counter-current flow reactor, wherethe direction of travel for the solid particles is the opposite of thedirection of travel for the fluid phases. U.S. Pat. Nos. 4,968,409 and5,916,529 provide examples of moving bed reactors designed forcounter-current flow. The reactors include a distributor thatcorresponds to a cone for guiding the catalyst particles into a pipe asthe catalyst moves down through the reactor. The cone distributorincludes openings to allow gas to pass through the cone. The conedistributor also includes liquid conduits to transfer fluid from areservoir up to the catalyst in the cone distributor. While acounter-current flow reactor can handle a three-phase flow, managing thethree-phase flow is difficult. For example, the flow rates for eachphase need to be balanced to avoid flooding of the reactor.Additionally, the residence time for contact between the liquid and thecatalyst particles is relatively high, so reactions requiring a shortcontact time between the liquid and the solid phases are not suitablefor this type of counter-current reactor configuration.

European Patent Application EP 0552457 describes another example of acounter-current moving bed reactor configuration.

U.S. Pat. Nos. 7,371,915 and 7,414,167 describe co-current moving bedreactor systems for conversion of oxygenates to propylene. Because theconversion reaction converts low molecular weight oxygenates topropylene, liquid is not formed in the reactors.

U.S. Pat. No. 8,323,476 describes moving bed hydroprocessing reactorsfor hydroprocessing of liquid feeds. The amount of hydrogen introducedinto the reactors is limited so that a continuous liquid phase ismaintained within the hydroprocessing reactors. The liquid is contactedwith the solids in a radial flow configuration.

U.S. Pat. No. 5,849,976 describes a moving bed solid catalysthydrocarbon alkylation process. The reaction zone is operated at liquidphase conditions.

U.S. Pat. No. 9,162,205 describes a co-current moving bed reactor systemfor contacting fluids with adsorbent particles. Due to the nature of anadsorbent/desorbent system, maldistribution of fluid flow within thereactor may lead to reduced performance, but does not otherwise resultin problems due to excessive reaction of fluids with catalyst particles.

U.S. Patent Application Publication 2017/0121237 describes a two reactorsystem for conversion of oxygenates to naphtha and distillate. A firstreactor is used to convert the oxygenates to a mixture that includesolefins. After passing through a separation stage, at least a portion ofthe olefins are then passed into a second reactor for oligomerization ofolefins to form naphtha and distillate boiling range compounds.

U.S. Patent Application Publication 2017/0137342 describes multi-phaseseparators for use in producing oxygenates and olefins fromhydrocarbons. The multi-phase separators are described as being suitablefor use in moving bed reactors.

What is needed are systems and methods to enable transfer of aco-current three-phase flow from one moving bed reactor to anothermoving bed reactor when performing reactions, such conversion ofoxygenates to distillate boiling range compounds, where it is desirableto control contact time of fluids with catalyst while also managing flowuniformity. This can include having the ability to separate athree-phase flow so that each phase can be separately introduced at acontrolled rate. This can further include introducing each phase in amanner that results in substantially uniform mixing of the phases.

SUMMARY

In an aspect, a method for upgrading a feed to form distillate boilingrange compounds is provided. The method can include passing a catalystflow comprising a catalyst into a first moving bed reactor of aplurality of serially connected moving bed reactors. The first movingbed reactor can include a first reactor feed inlet and a first reactoreffluent outlet. The catalyst can include at least one of oxygenateconversion catalyst and olefin oligomerization catalyst. The method canfurther include exposing a feed to the catalyst flow in the first movingbed reactor under first reaction conditions to form a first partiallyreacted effluent. The first reaction conditions can include at least oneof oxygenate conversion conditions and oligomerization conditions. Atemperature differential between the first reactor feed inlet and thefirst reactor effluent outlet can be 85° C. or less. The method canfurther include stripping the catalyst flow with a first stripping fluidto separate at least a portion of the first partially reacted effluentfrom the catalyst flow. The at least a portion of the first partiallyreacted effluent can include a liquid phase effluent portion and a vaporphase effluent portion. The method can further include passing thestripped catalyst flow into a second moving bed reactor of the pluralityof serially connected moving bed reactors. The method can furtherinclude passing the liquid phase effluent portion into the second movingbed reactor as a substantially axial flow. Additionally, the method canfurther include exposing the vapor phase effluent portion to thestripped catalyst flow in the presence of the liquid effluent portion inthe second moving bed reactor under second reaction conditions to form asecond effluent including distillate boiling range compounds. The secondreaction conditions can include at least one of oxygenate conversionconditions and oligomerization conditions.

In some aspects, the feed can include at least a portion of an effluentfrom a third moving bed reactor of the plurality of serially connectedmoving bed reactors.

In some aspects where the catalyst comprises olefin oligomerizationcatalyst and where the first reaction conditions compriseoligomerization conditions, the method can further include converting anoxygenate feedstock in one or more reactors to form a conversioneffluent comprising a conversion total hydrocarbon product, theconversion total hydrocarbon product comprising 20 wt % or more olefins.In such aspects, the feed can include at least a portion of theconversion total hydrocarbon product.

In another aspect, a system for upgrading oxygenates is provided. Thesystem can include a first plurality of serially connected moving bedreactors. Each moving bed reactor in the first plurality of seriallyconnected moving bed reactors can include a catalyst inlet, a feedinlet, a stripping fluid inlet, a catalyst outlet, and a vapor phaseeffluent outlet. The system can further include a separation stagecomprising a separation stage inlet and one or more separation stageoutlets, the separation stage inlet being in fluid communication with atleast at least one vapor phase effluent outlet. The system can furtherinclude a second plurality of serially connected second moving bedreactors. Each moving bed reactor in the second plurality of seriallyconnected moving bed reactors can include a second catalyst inlet, asecond feed inlet, a second stripping fluid inlet, a second catalystoutlet, a second liquid phase effluent outlet, and a second vapor phaseeffluent outlet. At least one second feed inlet can be in fluidcommunication with at least one of the one or more separation stageoutlets. The system can further include a second separation stagecomprising a second separation stage inlet and one or more secondseparation stage outlets, the separation stage inlet being in fluidcommunication with at least one second liquid phase effluent outlet andat least one second vapor phase effluent outlet. Additionally, thesystem can further include a regenerator comprising a first regeneratorinlet, a first regenerator outlet, a second regenerator inlet, a secondregenerator outlet, an oxygen inlet, and a flue gas outlet. The firstregenerator inlet can be in solids flow communication with at least onecatalyst outlet and/or the second regenerator inlet being in solids flowcommunication with at least one second catalyst outlet. Additionally oralternately, the first regenerator outlet can be in solids flowcommunication with at least one catalyst inlet and/or the secondregenerator outlet can be in solids flow communication with at least onesecond catalyst inlet.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 shows a side view of an example of a feed distribution apparatus.

FIG. 2 shows additional details for portions of the feed distributionapparatus in FIG. 1.

FIG. 3 shows a bottom view of the feed distribution apparatus shown inFIG. 1.

FIG. 4 shows a side view an example of a moving bed reactor.

FIG. 5 shows a top view of the moving bed reactor shown in FIG. 4.

FIG. 6 shows a vertical stack configuration of serially connected movingbed reactors.

FIG. 7 shows a horizontal configuration of serially connected moving bedreactors.

FIG. 8 shows a configuration with separate series of moving bed reactorsfor performing oxygenate conversion and olefin oligomerization.

DETAILED DESCRIPTION

All numerical values within the detailed description and the claimsherein are modified by “about” or “approximately” the indicated value,and take into account experimental error and variations that would beexpected by a person having ordinary skill in the art.

Overview

In various aspects, systems and methods are provided for conversion ofoxygenate feeds to distillate boiling range products using multiplemoving bed reactor stages. The systems and methods allow for multiplestages to be used while avoiding the need for distillation or otherboiling point based separation as the mixture of feed and effluent ispassed between stages. Instead, a stripping gas is used to disengage thefeed and effluent from the catalyst solids. In combination with animproved moving bed reactor design, this can allow substantially all ofthe feed and effluent from a first moving bed reactor stage to be passedinto a second moving bed reactor stage, even when the feed and effluentinclude both vapor and liquid phase portions.

In some aspects, multiple moving bed reactors can be arranged into twogroups of stages. A first group of stages can be operated underconditions that favor oxygenate conversion, while a second group ofstages can be operated under conditions that favor olefinoligomerization. In such aspects, an intermediate separation can beperformed between the stages for oxygenate conversion and the stages forproducing distillate boiling range compounds by oligomerization. Inother aspects, the multiple moving bed reactor stages can be operatedunder a single set of conditions, so that the distillate product isformed without performing an intermediate distillation.

Upgrading of oxygenates to distillate boiling range products is an areaof continuing interest, based on the potential for oxygenate conversionto serve as a bridge between natural gas feedstocks (and/or other feedscontaining primarily methane and other C⁴⁻ alkanes) and higher valueproducts, such as distillate fuels and various specialty chemicals.Hydrocarbon reforming is one of the few practical options for conversionof small alkanes. Reforming generates synthesis gas, which can readilybe converted to methanol and/or other small alcohols. The methanol(and/or other small alcohols) formed in this manner can then be used asa feedstock for processes that allow for upgrading of oxygenates.

Unfortunately, there are a number of challenges to performing oxygenateconversion to distillate boiling range products at a commercial scale.Some difficulties can be related to the highly exothermic nature ofupgrading oxygenates to distillate boiling range compounds. Bothconversion of oxygenates to olefins and olefin oligomerization arestrongly exothermic reactions. Due to practical considerations, it isdesirable to limit the temperature gradient across an oxygenateconversion reactor to roughly 85° C. or less (˜150° F. or less). Becauseof the exothermic nature of oxygenate conversion and/or olefinoligomerization, the limit on temperature gradient across a reactorconstrains the size of a reactor and/or the space velocity within areactor. As a result, managing heat during conversion of oxygenates andoligomerization of olefins generated from oxygenate conversion can posesignificant challenges in a fixed bed reactor environment.

Using a plurality of moving bed reactors can mitigate or minimize suchheat management difficulties. For example, a plurality of moving bedreactors can be used to perform the conversion and oligomerizationreactions. The reactors can be sized and/or operating conditions can beselected so that the amount of temperature increase across a singlereactor is less than a target value, such as having a temperature riseof 85° C. (˜150° F.) or less across a reactor, or 75° C. or less, or 60°C. or less. An initial feed, which may only correspond to a gas phasefeed of oxygenates and light olefins, can then be passed into the firstreactor. The plurality of moving bed reactors can then be used tofacilitate substantially complete conversion of oxygenates as well asoligomerization of the resulting olefins to a desired degree. This canallow, for example, for conversion of an oxygenate feed (optionally alsoincluding light olefins) to distillate boiling range products whileavoiding excessive heating within any single reactor.

Other difficulties can be related to maintaining catalyst activityduring upgrading of oxygenates to distillate boiling range compounds.Without being bound by any particular theory, it is believed that thereare two primary modes of catalyst deactivation for the zeotype catalystsused in oxygenate conversion and olefin oligomerization. One type ofdeactivation is due to coke formation on the catalyst. As cokeaccumulates, it is believed that active sites can be blocked, resultingin lower catalyst activity. Fortunately, such coke can be removed byregeneration at high temperature, which can restore a substantialportion (such as up to all) of the activity loss due to coke formation.

In a fixed bed catalyst environment, removal of coke from catalyst canonly occur during dedicated regeneration periods. In between suchregeneration periods, the primary option for extending run length in afixed bed reactor to a commercially desirable level is to build a largercatalyst bed. At the start of an oxygenate conversion process, thecatalyst near the top of the bed performs most of the oxygenateconversion. This results in coking of the catalyst. As cokingdeactivates the catalyst near the top of the bed, catalyst at locationsfarther down in the bed performs an increasing percentage of theoxygenate conversion. Thus, the depth of the fixed bed can be increasedso that the run length between regeneration steps is at a desirablelevel.

Although creating a larger fixed bed can be effective for overcomingdifficulties due to coke formation, such increases in the size of afixed bed can actually increase another type of deactivation. Withoutbeing bound by any particular theory, it is believed that a second typeof catalyst deactivation can be due to steaming of the catalyst, or inother words exposing the catalyst to water at elevated temperatures.

In addition to forming olefins, the oxygenate conversion reactiongenerates a substantial amount of water. For example, if methanol isused as the oxygenate feed, two moles of water are generated for eachmole of ethene produced by the oxygenate conversion reaction. In a fixedbed environment, at the beginning of an oxygenate conversion reaction,the catalyst near the top of the catalyst bed can perform substantiallyall of the oxygenate conversion. This results in creation of water,which then is exposed to the remaining portion of the catalyst bed asthe feed and effluent passed through the bed toward the reactor exit. Asthe depth of a catalyst bed is increased to provide longer run lengthbetween regeneration, a corresponding increase occurs in the amount ofsteaming that is performed on catalyst near the bottom of the bed. Thus,by the time that catalyst near the bottom of the bed starts toparticipate in oxygenate conversion, such catalyst can undergosubstantial deactivation due to catalyst steaming.

Using a plurality of moving bed reactors can reduce or minimize theimpact of both of the above types of catalyst deactivation. With regardto deactivation due to coking, the size of a moving bed reactor can beselected relative to the expected velocity of catalyst within the movingbed, so that the catalyst can be regenerated with a desired frequency.This can maintain coke on catalyst at less than a target level. Withregard to deactivation due to steaming, using a moving bed reactorsystem means that only the catalyst currently participating in aconversion or oligomerization reaction is exposed to steam. When notinside a moving bed reactor, the catalyst can be disengaged from theliquid phase and gas phase portions of the flow. Additionally, catalystcan be replaced at a convenient rate, so that the average steamingexposure of the catalyst is less than a target value. Thus, using aplurality of moving bed reactors can both reduce the amount of catalystexposure to steam relative to the amount of feed processed, and can alsoallow for control over the average steam exposure prior to replacementof the catalyst particles.

In various aspects, upgrading of oxygenates in a multi-stage moving bedreactor system can be facilitated by using moving bed reactors thatinclude a feed distribution apparatus that is suitable for introducing a3-phase flow under co-current flow conditions. The feed distributionapparatus can allow for separate introduction of liquid and solids in amanner that allows for even distribution of liquid within the solids.The gas portion of the flow can be introduced in any of a variety ofconvenient manners for distributing gas into a liquid or solid flow.

The distribution apparatus allows for efficient and/or substantiallyeven distribution of a co-current axial liquid flow in a solid particleflow based on the relative angle of introduction for the liquid and thesolid particles. The solid particles can be introduced into the reactorby allowing the particles to drop under gravitational pull. The conduitdropping the particles can also be narrower than the portion of thereactor that is receiving the particles. This can result in the solidparticles forming a cone based on the angle of repose for the solidparticles. The liquid can then be introduced at a plurality of locationsaround the cone. The distribution channels for introducing the liquidcan be angled at the exit point, so that the liquid has a lateralvelocity component. Introducing the liquid with a lateral velocitycomponent can facilitate mixing of the liquid with the solid particles.However, even though the liquid initially has a lateral velocitycomponent, at least a majority of the liquid travels axially with thecatalyst through substantially the full length of the reactor prior todisengagement of the majority of the liquid from the solids. Any gas ina three-phase flow can be contacted with the solid and liquid in anaxial flow manner, a radial flow manner, or in any other convenientmanner that allows for a desired distribution pattern.

Because the solid particles are being dropped into the reactor to allowformation of a cone, several options are available for controlling therate of catalyst flow through the moving bed reactors. In some aspects,the flow rate of catalyst particles through a moving bed reactor can becontrolled using a catalyst flow controller valve at the bottom of thereactor. This type of control over the catalyst flow through a movingbed reactor can be used in any convenient configuration one or moremoving bed reactors. Using a catalyst flow controller valve at thebottom of a reactor can also limit the flow of catalyst delivered to thenext moving bed reactor. Limiting the catalyst flow rate into a reactorcan ensure that sufficient space is available in the reactor volume forthe desired cone to form at the angle of repose. In addition to thecatalyst flow controllers at the bottom of each reactor, a catalyst flowcontroller can also be included between the catalyst source and thefirst reactor in a series of reactors. For example, a flow controllercan be placed at the exit of a catalyst regenerator that is used toprovide regenerated catalyst to the first reactor in a series ofreactors. An example of a suitable flow controller is described in U.S.Pat. No. 10,188,998.

In other aspects where a plurality of moving bed reactors are arrangedas a vertical stack, the flow rate to all of the reactors in thevertical stack can be controlled using two catalyst flow controllers,plus the geometry of the transfer conduit(s) between reactors. In thistype of configuration, a substantial portion of the flow rate controlcan be dependent on gravity-assisted flow control. One of the flowcontrollers can be used to control the catalyst flow rate into the firstor top reactor in the vertical stack of reactors. A second flowcontroller can be used at the bottom of the final reactor in thevertical stack. This can control the overall flow of catalyst throughthe series of reactors. The geometry of the conduits between thereactors can then be used to control the flow rate of catalyst betweenreactors. In particular, the conduit size between reactors can narrow torestrict flow of catalyst between reactors. This can allow sufficienthead space to be available at the top of each catalyst bed so that acone of catalyst can form at the angle of repose.

It is noted that the distributor apparatus can work in conjunction withmethods for separating a three-phase flow as the flow exits from themoving bed reactor. The flow exiting a moving bed reactor can correspondto a oxygenate conversion effluent, an oligomerization effluent, or anoxygenate upgrading effluent (when a single series of reactors is usedfor both oxygenate conversion and olefin oligomerization). By separatingthe three-phase flow into gas, liquid, and solid components, thecomponents can be re-combined in a subsequent moving bed stage using thedistributor apparatus. The separation of the fluid phases from thecatalyst flow can be effective for separating 95 mol % or more of thehydrocarbons in the oxygenate upgrading effluent/conversioneffluent/oligomerization effluent from the catalyst flow.

One example of a suitable method for separating the three-phase flowinto gas, liquid, and solid components can be to use a stripping gas incombination with concentric pipes to allow for separate capture of thegas and liquid. For example, the stripping gas can be passed through thesolid particles at a location prior to the solids exit conduit. This cancause any liquids and gases entrained with the solid particles to bedriven out of the volume containing the solids and into a separatevolume, such as an outer pipe of a pair of concentric pipes. The wallbetween the inner pipe and the outer pipe can include protectedopenings, such as bubble caps, that allow transport of gas from theouter pipe to the inner pipe while minimizing transport of liquids. Theliquids can instead accumulate at the bottom of the outer pipe and exitfrom openings that can be accessed when the accumulated liquid level issufficiently high. U.S. Patent Application Publication 2017/0137342shows an example of this type of structure.

In this discussion, the “solids volume” within a reactor is defined asthe volume that receives solid catalyst particles to form the movingbed. In various aspects, the solid particles are introduced at or nearthe top of the solids volume, and form a cone at the angle of repose forthe solid particles. The solids volume includes the solids exit volume,where the mixture of solids, liquid, and any remaining gas are strippedfrom the solids using a stripping gas. In some aspects, the bottom ofthe solids exit volume corresponds to the bottom of the solids volume.In other aspects, the bottom of the solids volume can correspond to thebottom of the exit port(s) for the stripping gas used in the solids exitvolume.

In this discussion, the “reaction zone volume” corresponds to a regionwithin the solids volume. The top of the reaction zone volumecorresponds to the base of the cone that forms at the angle of repose inthe solids volume. The bottom of the reaction zone volume corresponds tothe beginning or top of the solids exit volume, where the solids arecontacted with stripping gas. The top of the solids exit volume can bedefined based on a change in the geometry, such as the transition from acylinder or annular shape to a cone shape, or the top of the solids exitvolume can correspond to the top of the exit port(s) for the strippinggas used in the solids exit volume.

In this discussion, operating a reactor to have a majority of the liquidtravel axially with the solid particles can be characterized based onone or more of the following features. In some aspects, 40 vol % or more(or 50 vol % or more) of the liquid that contacts the solid particles inthe reactor can be initially brought into contact with the solidparticles in the top 20% of the volume occupied by the solid particles,such as up to substantially all of the liquid. In other words,regardless of the length of the contact time with the particles, theinitial contact can be in the top 20% of the volume occupied by thesolid particles. In many aspects, this will have substantial overlapwith the top 20% of the solids volume, but the top 20% of the volumeoccupied by the solid particles can differ from the top 20% of thereaction zone volume in the reactor if there is substantial distancebetween the top level of the solid particles and the top of the reactor.By definition, any liquid that first comes into contact with a topsurface of the catalyst bed in the moving bed reactor corresponds toliquid that first contacts the top 20% of the volume occupied by thesolid particles.

Additionally or alternately, in some aspects 40 vol % or more (or 50 vol% or more) of the liquid that contacts the solid particles can beseparated from the solid particles in the bottom 20% of the volumeoccupied by the solid particles, such as up to substantially all of theliquid. In many aspects, this will have substantial overlap with thebottom 20% of the reactor volume, but the bottom 20% of the volumeoccupied by the solid particles can differ from the bottom 20% of thevolume in the reactor if there is substantial distance between thebottom level of the solid particles and the bottom of the reactorvolume.

It is noted that “top” and “bottom” are relative to the direction of theco-current flow of liquid and solid particles within the reactor. Invarious aspects, it can be convenient to align the direction of flowwith the direction of gravitational force, in order to reduce orminimize maldistribution of liquid relative to the solid particles dueto gravitational pull. However, if a reactor is oriented in anothermanner, the “top” and “bottom” of the solid particle bed can be definedso that the “top” corresponds to where solid particles are added to thebed and the “bottom” corresponds to where solid particles are removedfrom the bed (such as by exiting the reactor and passing into a transferpipe). It is noted that in an upflow configuration, this would result inthe “top” of the moving bed being closer to the bottom of the reactor,while the “bottom” of the moving bed would be closer to the top of thereactor.

In this discussion, operating a moving bed reactor with a three-phaseflow corresponds to operating a reactor where 45 vol %-70 vol %, or 50vol % to 70 vol % of the reaction zone volume corresponds to a solid(particles) phase; 10 vol % or more of the reactor volume corresponds toa liquid phase, such as 10 vol % to 45 vol %, or 20 vol % to 45 vol %,or 10 vol % to 35 vol %, or 20 vol % to 35 vol %, or 10 vol % to 30 vol%, or 10 vol % to 25 vol %; and 5 vol % or more of the reactor volumecorresponds to a gas phase, such as 5 vol % to 40 vol %, or 10 vol % to40 vol %, or 5 vol % to 35 vol %, or 5 vol % to 30 vol %, or 10 vol % to30 vol %, or 5 vol % to 25 vol %, or 5 vol % to 20 vol %.

In this discussion, fluid communication is defined as the ability forvapor and liquid to move between two process elements. Vaporcommunication is defined as the ability for vapor to move between twoprocess elements, while having reduced, limited, or optionally nomovement of liquid between such process elements. Solids flowcommunication is defined as the ability for solid particles to movebetween two process elements, which typically means that fluidcommunication is also possible.

Serial Moving Bed Reactors for Upgrading Oxygenates

In various aspects, a plurality of moving bed reactors, arranged forserial processing, can be used for upgrading of oxygenates to formdistillate boiling range compounds. The plurality of reactors can be anyconvenient number that allows conversion of the feed to reach a desiredlevel of completion. In some aspects, the number of reactors can beselected to allow for substantially complete conversion of an oxygenatefeed to olefins. In other aspects, the number of reactors can beselected to achieve a desired level of oligomerization (such as adesired yield of distillate boiling range compounds) for the olefinsformed during oxygenate conversion.

Some considerations for selection of reactor size and number of reactorscan be related to managing the temperature gradient across each reactor.Preferably, the temperature gradient across a single reactor can be 80°C. or less (˜150° F. or less), or 70° C. or less, or 60° C. or less. Duein part to higher feed concentration of oxygenates at the beginning of aseries of reactors, if a series of reactors includes equal amounts ofcatalyst (i.e., reactors of equal size), the temperature gradient acrossthe first reactor in the series, and then decrease across eachsuccessive reactor. For example, for three reactors in series, thetemperature gradient across the first reactor would be larger than thetemperature gradient across the second reactor, while the second reactortemperature gradient would be larger than the third reactor temperaturegradient. In order to account for this, the size/amount of catalyst forreactors later in the series can be larger than reactors earlier in theseries. This can increase the time of catalyst exposure to feed in thelater reactors. In some aspects, this can be used to allow two or morereactors in series to have a similar temperature gradient, such as up toall reactors in a series. In other aspects, the first reactor in aseries can have a different, smaller size/smaller amount of catalystthan at least one other reactor in the series, with each reactor in theseries being at least as large/containing at least as much catalyst asthe prior reactor. Still other potential methods for selecting reactorsize/catalyst amount can also be used. It is noted that catalyst amountrefers to the amount of catalyst used during operation of a moving bedreactor. Thus, in configurations where a separate catalyst flowcontroller is used after each reactor, it can be possible to have movingbed reactors with different amounts of catalyst even though the volumefor holding catalyst in each reactor is similar.

With regard to managing temperature, during or after removal of catalystand fluids from a reactor, one or more steps can be taken to reduce thetemperature in the next reactor. For example, the stripping gas used todisengage fluids from the catalyst can also serve as a quench gas. Thiscan allow the inlet temperature two or more reactors in a series, suchas up to each reactor in a series, to be substantially similar. Reactorscan be considered to have substantially similar inlet temperatures ifthe average temperature of catalyst and fluid entering the reactor iswithin 10° C. The average temperature of catalyst and fluids entering areactor can correspond to a weighted average based on the weight of eachcomponent entering the reactor during a unit time.

Other considerations for selection of reactor size and/or number ofreactors can be related to the desired product for a series of reactors.In some aspects, both oxygenate conversion and olefin oligomerizationcan be performed in a series of reactors. In such aspects, the spacevelocity and other reaction conditions can be selected to balanceconversion of oxygenates within the feed with oligomerization of theresulting olefins. In other aspects, two series of reactors can be usedso that separate conditions can be provided for oxygenate conversion andolefin oligomerization. For example, even though the same zeotypecatalyst can be used for both oxygenate conversion and olefinoligomerization, favorable conditions for olefin oligomerization tend tocorrespond to lower temperatures and higher pressures relative tofavorable conditions for oxygenate conversion. In such aspects, reactorsizes and conditions in a first series of reactors can be selected toprovide substantially complete conversion of oxygenates, while reactorsizes and conditions in a second series of reactors can be selected toprovide a desired level of oligomerization.

Configuration Example 1—Vertically Stacked Reactors

One option for arranging reactors in series can be to use a verticalstack of reactors. This can allow gravity-assisted flow to be used tomanage the flow of catalyst between reactors.

FIG. 6 shows an example of a series of moving bed reactors arranged as avertical stack. In a configuration such as FIG. 6, the reactors in thevertical stack can correspond to reactors R1-Rn, where n is greater thanor equal to 2, with R1 corresponding to the initial entry point for feedand Rn corresponding to the final reactor in the series.

In FIG. 6, an oxygenate feed 600 such as methanol, dimethyl ether, orboth, and a quench feed such as water or an inert gas, are introduced atthe top of the first stage reactor R1. A flow 710 of catalyst particlesenter the first stage reactor R1 at the top of the reactor. The reactorcould be either axial or radial flow with respect to flow of vapor inthe reactor. To the degree that a liquid portion is present in the feedin first stage reactor R1 or any subsequent stage, the liquid portion ofthe feed can substantially flow axially through the reactor. As thereaction takes place, the heat release from the exothermic dehydrationof the oxygenates (for conversion to olefins) increases the temperatureinside the reactor bed, and also deposits coke on the catalystparticles. The first stage reactor R1 contains the highest extent ofreaction caused by the reaction of the fresh feed over the activecatalyst. Therefore, in some aspects, the amount of catalyst included inreactor R1 can correspond to the smallest amount of catalyst in reactorsR1-Rn. At or near the bottom of R1, the unreacted feed and products 610are separated from the catalyst and removed using a stripping fluid 720such as water, an inert gas, or a portion of the tail gas 682 fromproduct separation unit 685. Optionally, additional feed 601 could alsobe injected as part of stripping fluid 720 if staged injection isdesired. After disengagement of fluids, the catalyst particles exit R1through a downcomer line 620 that connects R1 to R2. The stripping fluid720 could also act as the interstage quench. Similarly, interstage heatexchanger 851 can be used to further control the temperature of the 630stream before the inlet to R2. The second stage reactor R2 operates inthe same manner as the previous stage R1. As the reaction progresses,the catalyst particles continue to coke and deactivate. At the bottom ofR2 stripping fluid 730 is introduced to separate and remove theremaining unreacted feed and products 640 from the catalyst. Thestripping fluid could be water, inert gas, or portion of the tail gas683 from product separation unit 685. An interstage heat exchanger 852can be used to adjust the temperature of stream 660 before the inlet tothe next subsequent stage. A downcomer line 650 connects R2 to the nextstage reactor for the transfer of the catalyst. This process continuesall the way through the last stage Rn where full conversion of thereactant occurs. The final products are removed using a stripping fluidstream 740. The stripping fluid could be water, inert gas, or portion ofthe tail gas 684 from product separation unit 685. The resultingeffluent stream 670 is sent to a first separation unit block 685 whichcould be a flash tank or a distillation column. In the example shown inFIG. 6, the lighter gases 680 (such as tail gas) are separated from themain liquid products 690. All or, portion of the tail gas can berecycled back 681 for use in the feed to the reaction system, and/orrecycled 682, 683, and/or 684 for use as stripping fluid in one or morestages of the reaction system. The liquid product 690 from the firstseparation unit 685 can further be fractionated in a second separationunit 695 into a gasoline cut C⁹⁻ 770 and a distillate cut C₉₊ 780. Aportion or all of the gasoline cut 771 can be recycled back to theappropriate last stages of the reactor to improve the oligomerizationreaction and produce heavier molecular weight product (e.g., increasedistillate make). The products from the second separation unit can befurther fractionated and/or hydrotreated as needed.

The flow of the catalyst in the reactor is maintained using agravity-assisted catalyst flow controller valve 800 at the exit of thelast stage reactor Rn. The coked catalyst exits Rn and is conveyed 801to a regenerator unit 901 where air 750 is introduced at the bottom. Thedeposited coke on the catalyst particles is burnt with the oxygen in theair, resulting in generation of CO₂ and steam. The flue gas 760 from thecombustion is removed from the top of the regenerator unit. The catalystflow in the regenerator can be adjusted using a catalyst flow controllervalve 810. The regenerated catalyst from regenerator 901 is conveyedback 710 to the top of the first reactor stage R1 for the process tocontinue. A transfer mechanism (not shown) such as a conveyor belt, apneumatic conveyor, or a catalyst lift system using steam or other gasescan be used to transfer the catalyst particles from regenerator 901 toR1. Additionally, make up catalyst if needed can be added to stream 710.

Configuration Example 2—Horizontal Serial Reactors

FIG. 7 shows another example of a potential configuration for a seriesof reactors. In FIG. 7, the reactors are shown in a horizontalconfiguration, but any convenient arrangement of reactors (includingvertical) could be used for such a configuration. This is due in part tothe use of catalyst flow controllers between each moving bed reactorstage, so that the reaction mechanism and coke deactivation in thisexample is similar to the process described above and shown in FIG. 6.Similarly, since the extent of the reaction is the same, the design andsizes of the reactor stages follow a similar pattern.

In the configuration shown in FIG. 7, an oxygenate feed 600 (includingany quench feed) is introduced at the top of the first stage reactor R1.A flow 710 of catalyst particles enter the first stage reactor R1 at thetop of the reactor. The reactor could be either axial or radial flowwith respect to flow of vapor in the reactor. To the degree that aliquid portion is present in the feed in first stage reactor R1 or anysubsequent stage, the liquid portion of the feed can substantially flowaxially through the reactor. At the bottom of R1, the unreacted feed andproducts 610 are separated from the catalyst and removed using astripping gas or fluid 720. The stripping fluid could be water, inertgas, or portion of the tail gas 682 from product separation unit 685.Additional reactant 601 could also be injected at this location ifstaged injection is desired. The catalyst particles exit R1 through atransfer line 620 that connects R1 to R2. The flow of the catalyst outof the first stage reactor R1 is maintained using a catalyst flowcontroller valve 800. A transfer mechanism (not shown) such as aconveyor belt, a pneumatic conveyor, or a catalyst lift system usingsteam or other gases can be used to transfer the catalyst particles fromR1 to R2. The stripping gas 720 also acts as the interstage quench.Similarly, interstage heat exchanger 851 can be used to further controlthe temperature of the stream 630 before the inlet to R2. The secondstage reactor R2 operates in the same manner than the previous stage R1.As the reaction progresses, the catalyst particles continue to coke anddeactivate. At the bottom of R2 stripping gas/quench 730 is introducedto separate and remove the remaining unreacted feed and products 640.The stripping fluid could be water, inert gas, or portion of the tailgas 683 from product separation unit 685. An interstage heat exchanger852 can be used to adjust the temperature of stream 660 before the inletto the next subsequent stage. A catalyst transfer line 650 connects R2to the next stage reactor for the transport of the catalyst. Catalystflow controller valve 810 adjusts the flow of the particles in R2. Atransfer mechanism (not shown) such as a conveyor belt, a pneumaticconveyor, or a catalyst lift system using steam or other gases can beused to transfer the catalyst particles from R2 to the next reactorstage. This process continues all the way through the last stage Rnwhere full conversion of the reactant occurs. The final products areremoved using a stripping gas/quench stream 740 and the effluent stream670 is sent to a recovery block. The stripping fluid could be water,inert gas, or portion of the tail gas 684 from product separation unit685. The effluent stream 670 is sent to a first separation unit block685 which could be a flash tank or a distillation column. The lightergases 680 (i.e., tail gas) are separated from the main liquid products690. All or a portion of the tail gas can be recycled back 681 for usein the feed to the reaction system, and/or recycled 682, 683, and/or 684for use as stripping fluid in one or more stages of the reaction system.The liquid product 690 from the first separation unit can further befractionated in a second separation unit 695 into a gasoline cut C⁹⁻ 770and a distillate cut C₉₊ 780. At least a portion of the gasoline cut 771can be recycled back to the appropriate last stages of the reactor toimprove the oligomerization reaction and produce heavier molecularweight product (e.g., increase distillate make). The products from thesecond separation unit can further be hydrotreated as needed. The flowof the catalyst in the last reactor stage is maintained using a catalystflow controller valve 820.

It is noted that in some aspects, the catalyst flow controllers 800,810, and 820 can be set to the same flow rate to reduce, minimize, orprevent accumulation or depletion in any bed(s). In other aspects, ahopper can be included on top of each bed that is specifically sized forthe corresponding flow rate of that bed. In such aspects, each flowcontroller can be set to a separate rate. For example, in a horizontalconfiguration including a hopper on top of each reactor, the flow rateof catalyst in each moving bed can be selected based on the rate ofdeactivation in the corresponding reactor.

The coked catalyst exits Rn and is conveyed 801 to a regenerator unit901 where air 750 is introduced at the bottom. The deposited coke on thecatalyst particles is burnt with the oxygen in the air, resulting ingeneration of CO₂ and steam. The flue gas from the combustion is removedfrom the top of the regenerator unit 760. The catalyst flow in theregenerator can be adjusted using a catalyst flow controller valve 830.The regenerated catalyst from regenerator 901 is conveyed back 710 tothe top of the first reactor stage R1 for the process to continue. Atransfer mechanism such as a conveyor belt, a pneumatic conveyor, or acatalyst lift system using steam or other gases can be used to transferthe catalyst particles from regenerator 901 to R1. Additionally, make upcatalyst if needed can be added to stream 710. In this type ofconfiguration, the rate of the catalyst flow from one stage to anothercan be independently adjusted based on performance and process criteria,hence providing an additional degree of freedom to the operation.

Configuration Example 3—Separation of Oxygenate Conversion andOligomerization Stages

FIG. 8 shows yet another configuration example for upgrading ofoxygenates to distillate boiling range compounds. In the configurationshown in FIG. 8, two separate series of reactors are used. One series ofreactors corresponds to a series of reactors for conversion ofoxygenates to olefins, while a second series of reactors receives atleast a portion of the conversion effluent as an olefinic feed foroligomerization to form distillate boiling range compounds. Optionally,both series of reactors can correspond to moving bed reactors. AlthoughFIG. 8 shows an arrangement of vertical stacks of reactors, anyconvenient arrangement can be used, such as a horizontal arrangement.

The oxygenate conversion reaction and the olefin oligomerizationreaction can be carried out over the same catalyst or differentcatalysts. In the former case when the same catalyst is employed in thetwo reactor systems, a common regenerator unit can be shared between thetwo as shown in FIG. 8. In this configuration, the oxygenate dehydrationreaction, e.g., methanol conversion to olefins, takes place in the firstreactor system and the conversion of the olefins to final hydrocarbonfuel and chemicals takes place in the second reactor system. In thelatter case where two different catalysts are used, two separateregeneration units can be used. Various potential configurations can beused when separate series of moving bed reactors are used for eachreaction. For example, the two reactor systems can have the same numberof stages or a different number of stages; the two reactor systems canbe both vertically or horizontally stacked, or one stack could bevertical while the other be horizontal; the two reactor systems can bothbe axial flow or radial flow (for flow of gas), or one reactor systemcould be radial flow for gas while the other is axial flow.

The configuration in FIG. 8 shows two multi-stage reactor systemsstacked vertically, corresponding to reactor series A (for oxygenateconversion) and reactor series B (for olefin oligomerization). The feed100 enters the first reactor stage A1 in the first stack. The operationin this reactor system is similar to the one described in FIG. 6. Theflow of the catalyst is controlled using the catalyst flow control valve910 at the bottom of the last reactor stage in the first stack. Theproducts from the first reactor stack can be sent to a first separationstage 960. The first separation stage 960 can fractionate the productsfrom the first reactor stack and send the intermediate products 680 tothe second reactor stack for further conversion. Any convenient typesand/or numbers of separators can be used in separation stage 960. Insome aspects, the water formed in the first step can be removed 690 inthis separation unit and the intermediate hydrocarbon products such aslight olefins can be sent as feed 680 to the second stack reactorsystem. Optionally, a portion of the product gas 680 can be recycledback 681 for use in the feed to the reaction system, and/or recycled682, 683, and/or 684 for use as a stripping fluid for one or more stagesin the reaction system. The second reactor system has N reactor stagesstacked vertically, B1 to Bn. In the configuration shown in FIG. 8, feed680 enters the first reactor stage B1. Fresh and/or regenerated catalyst850 also enters reactor stage B1 at the top. The products from eachreactor stage (750 and 780) are removed at the bottom using a strippingfluid (860 and 870). The stripping fluid could be water, inert gas, orportion of the tail gas 321 or 322 from the second product separationunit 985. The catalyst is transferred to downstream stages through adowncomer pipe (760 and 790). The removed products and unreacted feedfrom each stage are further cooled in external heat exchangers (853 and854). In the last reactor stage Bn, the final product 810 is removedusing a stripping fluid 880 and sent to the second product recovery unit985 where fractionation takes place. For example, fractionation canallow for separation into a tail gas 820 and liquid hydrocarbon products830. The stripping fluid 880 could be water, inert gas, or portion ofthe tail gas 823 from the second product separation unit 985. Moregenerally, the tail gas from second separation unit can be recycled 821,822, and/or 823 for use as stripping fluid in on or more stages of thereaction system. The liquid hydrocarbon product 830 from the secondseparation unit can further be fractionated in a third separation unit995 into a gasoline cut C⁹⁻ 910 and a distillate cut C₉₊ 920.Optionally, at least a portion of the gasoline cut 911 can be recycledback to the appropriate last stages of the reaction system to improvethe oligomerization reaction and produce heavier molecular weightproduct (e.g., increase distillate make). The products from the secondseparation unit 985 can further be hydrotreated as needed. The finalproducts 920 can further be hydrotreated treated.

The catalyst flow in the second reactor stack is controlled using acatalyst flow controller 930 at the bottom of the last reactor stage Bn.The spent catalysts from both reactor systems are sent to a commonregenerator unit 1001. A conveying mechanism system similar to the onedescribed above can be used to transfer the catalysts to the regeneratorunit 1001 through conduits 801 and 841 for first and second stackreactor systems respectively. The regenerator works in the same mannerdescribed above where air 890 enters at the bottom and the flue gas 1000exits at the top.

The catalyst flow at the bottom of the regenerator unit can becontrolled with one or two catalyst flow controller. The decision willbe based on the difference in the rate of catalyst deactivation betweenthe first reactor stack and second reactor stack. If both rates in bothreactor stacks are similar, then one catalyst flow controller can besufficient to distribute back substantially the same catalyst flow toboth reactor stacks. If the catalyst deactivation rates between thereactor stacks are different, catalyst flow controller 920 can beadjusted to provide the rate of catalyst to the first reactor stackthrough line 710 and catalyst flow controller 921 can be adjusted toprovide the rate of catalyst to the second reactor stack through line850.

Configuration Example 4—Example of Distributor Apparatus

In various aspects, upgrading of oxygenates to distillates in moving bedreactors can be performed using reactors that can provide relativelyuniform distribution of liquid in the catalyst particles within a movingbed reactor. One option for achieving a relatively uniform distributionis to use a distributor apparatus that can allow uniform distribution ofliquid in the catalyst particles under co-current axial flow conditionsfor the solids and liquids. When using such a distributor apparatus, thevapor flow in the moving bed reactors can correspond to radial flow oraxial flow.

FIG. 1 shows an example configuration for a distributor apparatus for amoving bed reactor. The distributor apparatus can be used for moving bedreactors where solids and liquids substantially travel through thereactor as an axial flow, while the gas flow can correspond to an axialflow, a radial flow, or any other convenient flow pattern. In someaspects, the distributor apparatus shown in FIG. 1 can be used inconjunction with a moving bed reactor where the catalyst is introducedalong the central axis of the reactor. This can result in the catalystfilling a central volume of the reactor. Such a configuration for amoving bed reactor can be suitable, for example, for a moving bedreactor where the catalyst, liquid, and gas all substantially traversethe reactor as axial flows. In other aspects, multiple instances of thedistributor apparatus in FIG. 1 can be arranged to provide catalyst foran annular catalyst volume in a moving bed reactor. This can bebeneficial for configurations such as the example shown in FIG. 4, wherethe catalyst and liquid in the moving bed reactor travel in asubstantially axial direction, while the gas in the feed contacts thecatalyst as a radial flow.

FIG. 1 shows a side view of an example of a distributor apparatus fordistribution of liquid when the solids are introduced along a centralaxis. In FIG. 1, a liquid distributor plate 101 is shown and highlightedwith hatch lines. The liquid distributor plate 101 fits into a movingbed reactor system through a system of connecting parts at the top andbottom of the distributor plate 101. The inlet feed pipes 102 areattached on the side of the top connecting part 103. Depending on theconfiguration, inlet feed pipes 102 can transport feed into the reactorin the form of gas, liquid, or a combination of gas and liquid. Forexample, in aspects where the gas portion of a feed contacts thecatalyst as a substantially radial flow, inlet feed pipes 102 cantransport a liquid portion of the feed into the reactor, withsubstantially all of the gas entering the reactor as a separate flow.

In FIG. 1, the catalyst particles enter the vessel through a solidsinlet conduit, such as catalyst feed line 104, which is attached in thecenter of the top connecting part 103. The distributor plate 101 sitsbelow the top connecting part 103. The portion of the feed provided byinlet pipes 102 drops into the distributor plate which includes one ormore concave shapes 105 with a radius to R (shown in FIG. 2). There area number of slots (or orifices) 106 placed around the distributor plate.Preferably, the slots can be placed evenly and concentrically around thedistributor plate. The slots provide fluid communication between theconcave shape(s) 105 and a solids volume 108. As shown in FIG. 2, theseslots have a depth at the top 261 of the slots 106 and a depth at thebottom 262 of the slots 106. The slots can optionally be threaded onboth ends (i.e., at top 261 and at bottom 262) to allow installation ofnozzles at both the top and the bottom (not shown). The nozzles cancorrespond to, for example, hollow cylinders with an opening at the topto allow the passage of the gas. In some aspects, the nozzles canfurther include a slit around the top side of the nozzle. This can allowliquid to flow through once a certain liquid height is built. This willprevent selective distribution of the liquid through certain nozzles andallow an even flow of the liquid through all the nozzles. It is notedthat the nozzles at the bottom 262 of the slots 106 can have a geometrythat is selected to facilitate distribution of the expected type of feedthat is passing through the slots 106, such as nozzles selected fordistribution of gas, liquid, or a mixture of gas and liquid.

Depending on the aspect, solids volume 108 can correspond to a singlecentral volume, an annular volume, or another convenient volume thatallows for axial flow of catalyst and liquid through a moving bedreactor. In aspects where solids volume 108 corresponds to a centralvolume in a reactor, a single solids inlet conduit 104 can providecatalyst to the solids volume 108. As another example, in aspects wheresolids volume 108 corresponds to an annular volume, a plurality ofdistributor apparatus can be arranged in a substantially symmetricmanner around the solids volume 108. Such a configuration is shown inFIG. 4.

In the example of a distributor apparatus shown in FIGS. 1 and 2, thebottom of the distributor plate 101 curves toward the inside of thesolids volume at a 45 degree angle (or more generally an angle between30° and 55°). The gas and/or liquid will enter the nozzles inserted intothe distributor plate's slots (or orifices) 106 and first drop adistance prior to then curving inward, such as at angle 272, whichcorresponds to a 60 degree angle in FIG. 2 (or more generally an anglebetween 45° and 70°). The gas and/or liquid can then travel a furtherdistance to exit through another (bottom) set of nozzles inserted intoslots 106, as shown in FIG. 2. As shown in FIG. 1, the distributor platesits on top of the bottom connecting part 107 which has a form of afunnel whose bottom diameter is that of the solids volume 108 and topdiameter is that of the distributor plate. The top connecting part 103,the distributor plate 101, and the bottom connecting part 107 are heldtogether with a bolted clamp 109. The bottom connecting part 107 caneither be welded to the top of the reactor vessel 108, as shown in FIG.1, or bolted with a flange (not shown).

At the interface 120 between catalyst feed line (or other solids inletconduit) 104 and solids volume 108, a characteristic width of thecatalyst feed line 104 can be smaller than a characteristic width of thesolids volume that is receiving the catalyst. The characteristic widthof the catalyst feed line 104 corresponds to the longest straight linethat can be drawn between two points on the catalyst feed line at theinterface 120 with the solids volume 108. In aspects when the catalystfeed line is roughly cylindrical, the characteristic width will be thediameter of the catalyst feed line. The width of the solids volumereceiving the catalyst particles can correspond to a) for a centralvolume, the diameter of a cylindrical volume as measured at the locationwhere the base of the catalyst cone forms in the reactor; b) for anannular volume, the radial distance between the outer surface and theinner surface that define the annular volume, at the location where thebase of the catalyst cone forms in the annular volume; or c) a similarlycharacteristic width for a volume having a shape other than an annularvolume or a cylindrical volume.

It is noted that the above definitions for the width of the solidsvolume are based on the location of where the base of the catalyst coneforms. Above the base of the cone, there can typically be a gap betweenthe interface of the catalyst feed line with the solids volume and thebase of the catalyst cone. This gap, which can be referred to as acontact zone or mixing zone, corresponding to the difference between thetop of the solids volume and the top of the reaction zone volume, allowsthe catalyst cone to form at the angle of repose.

During operation, the solid (catalyst) particles exit the catalyst feedline 104 and distribute inside the solids volume 108, forming a conicalshape 110 whose angle corresponds to the angle of repose of the solidparticles. The conical shape 110 is formed in part because thecharacteristic width of the solids inlet conduit(s) is smaller than thewidth of the volume receiving the solid particles. The angle of reposefor solid particles can vary, such as having an angle of repose ofroughly 10° to 40°. The bottom inner edge of the feed distributor platewhere it meets the catalyst feed line bottom can be slightly angled,such as having an angle of 17° for angle 271, as shown in FIG. 2. Moregenerally, angle 271 can range from 10° to 25°. The gas or gas/liquidexiting the distributor plate through exit surface 279, via the bottomnozzles of slots 106, can inject directly on top of the cone ofparticles (at the angle of repose) formed by the flow of solid particlesinto the reactor. This will allow enhanced fluid mixing from the top andlet the fluid evenly disperse radially as it flows downward. This is duein part to the lateral velocity of the feed toward the central axis,which can assist with having feed well-mixed with particles throughoutthe reaction zone volume.

In FIG. 2, the complement of angle 272 is angle 273. Angle 273corresponds to an angle that the exit surface 279 makes relative to aplane defined by interface 120 between the catalyst feed line 104 andthe reactor vessel 108. In FIG. 2, angle 273 is shown as having a valueof 30°, but more generally angle 273 (i.e., the angle of the exitsurface relative to the plane) can have a value between roughly 15° and45°.

FIG. 3 shows the bottom of the distributor plate. In this schematicthere are 12 orifice nozzles at the top 261 and bottom 262. The offsetbetween the center of the top and bottom nozzles as the distributorplate curves inward toward the center of the reactor vessel can also beseen in this figure. In various aspects, the number of the nozzles canbe determined based on the scale of the reactor vessel, the optimumdistance needed to space the nozzles, and the liquid mass flux.

It is noted that the above description contemplates having a distributorapparatus that is machined as a separate part from other parts of theoverall reaction system. In other aspects, the distributor apparatus canbe integrated with the overall reaction system in any convenient manner.For example, the distribution plate can be attached to the catalyst feedline (or other solids inlet conduit) 104, so that there is no visiblejoint between catalyst feed line 104 and the distributor plate. In suchan aspect, the “opening” in the distributor plate that allows catalystto pass from feed line (or other solids inlet conduit) 104 to the solidsvolume 108 can correspond to a part of the solids inlet conduit 104. Itis noted that the attachment between catalyst feed line 104 and thedistributor plate can correspond to a removable attachment, or thatattachment can correspond to the catalyst feed line 104 and thedistributor plate corresponding to a single piece. Alternatively, thedistributor apparatus and catalyst feed line 104 can be separate pieces,with the catalyst feed line 104 passing through an opening in thedistributor apparatus.

Configuration Example 5—Reactor with Annular Catalyst Volume (Radial GasFlow)

FIG. 4 shows an example of a moving bed reactor suitable for performinga co-current reaction in the presence of three phases. The reactor inFIG. 4 is designed to introduce the catalyst and liquid into an annularvolume. The solid and liquid can flow axially through the annularvolume. The gas phase portion of the feed can then be passed through thecatalyst and liquid in a substantially radial direction.

In FIG. 4, reaction vessel 401 includes an outer annular volume 402 andan inner annular volume 403. The inner annular volume 403 is the annularregion where the catalyst or other solid particles reside, and thereforecan be referred to as an annular solids volume. The outer annular volume402 and inner annular volume 403 are arranged around a central doublepipe or conduit, corresponding to an outer central pipe or conduit 404and an inner central pipe or conduit 405. The inner annular volume 403can include perforations (not shown) that permit vapor communicationbetween the inner annular volume 403 and outer central pipe 404. Theperforations can primarily allow gas to pass into the outer central pipe404, but some liquid can also pass through the perforations. Theperforations are small enough to retain substantially all of the solidparticles in the annular solids volume 403.

During operation, gas is introduced into the reactor 401 via a centralopening 406 which connects to an inlet pipe 407 with openings fordistributing the gas into outer annular volume 402. The tops of innerannular volume 403, outer central pipe 404, and inner central pipe 405are sealed at the top, so that the flow path for gas to reach the outercentral pipe 404 is by passing radially through inner annular volume403. During operation, solid particles (such as catalyst particles) areintroduced into the reactor 401 via a plurality of pipes 408. Theoutlets of the plurality of pipes 408 are roughly centered over themid-point of inner annular volume 403. Similar to the configurationshown in FIG. 1, the solid particles fall into the inner annular volume403 and form cones 409 corresponding to the angle of repose for theparticles. The liquid phase is passed into reactor 401 via a pluralityof liquid conduits 410. The liquid conduits 410 feed a plurality ofslots or openings 411 that are arranged around the pipes 408.Optionally, the slots or openings 411 can include nozzles. Also similarto FIG. 1, the slots or openings 411 are arranged to cause the liquidfeed to impinge on the cones 409, in order to facilitate evendistribution of liquid within annular volume 403. Optionally butpreferably, the nozzles 411 can be oriented so that the liquid exitingfrom slots or openings 411 has a lateral velocity component. In someaspects, pipes 408 and liquid conduits 410 can have a similarrelationship to inner annular volume 403 as the relationships betweencatalyst inlet flow 104, inlet pipes 102, and solids volume 108. In suchaspects, liquid conduits 410 can be in fluid communication with annularvolume 403 via the plurality of slots or openings 411, in a mannersimilar to how inlet pipes 102 are in fluid communication with solidsvolume 108 via slots or openings 106.

During operation, gas from outer annular volume 402 passes radially intoinner annular volume 403. This allows contact between gas, liquid, andsolid for performing a desired reaction. The gas then continues radiallyinto the first or outer central pipe 404. Outer central pipe 404includes a plurality of bubble caps 412 or other structures that canallow gas to pass through into inner central pipe 405 while retainingliquid entrained with the gas in outer central pipe 404.

At the bottom of reactor 401, the gas, liquid, and solids can beseparated to allow for further processing and/or or for introductioninto a subsequent moving bed stage. The gas exits through a main gasexit line 416 that is in fluid communication with the bottom of innercentral pipe 405. The solids can exit from inner annular volume 403 intoa plurality of solids exit volumes, such as cone-shaped exit volumes 414as shown in FIG. 4. A stripping gas 415 is passed through the solidsexit volumes 414 to strip liquid from the solids prior to allowing thesolids to exit via solids exit line 418. It is noted that a baffle 419connects the cones 414, so that the stripping gas cannot bypass thecones. The stripping gas causes liquid in the solid particles to exitinto the bottom of first or outer central pipe 404, where it is combinedwith any liquid collected by the bubble caps 412. The liquid canaccumulate at the bottom of outer pipe 404 to a sufficient height sothat the liquid can exit through openings 413 into liquid exit line 417.

FIG. 5 shows a top view of the reactor 401 shown in FIG. 4. In FIG. 5,the input conduits corresponding to central opening 406, pipes 408 (fortransfer of solid particles), and liquid conduits 410 are shown inrelation to each other. It is noted that liquid conduits 410 are usedwithin the reactor to provide liquid to a plurality of nozzles 411 (notvisible in FIG. 5) that are arranged around each pipe 408.

It is noted that the configuration shown in FIG. 5 includes a total of 8solids inlet conduits 408 and two liquid conduits 410. More generally,any convenient number of solids inlet conduits 408 and liquid conduits410 can be used, so long as liquid is distributed around each solidsinlet conduit. For example, in the configuration shown in FIG. 5, adistributor plate can be used to distribute the liquid feed from liquidconduits 410 around each of the solids inlet conduits 408. This can beperformed by a single distributor plate that includes all of the liquidconduits 410 and solids inlet conduits 408. Alternatively, a part ofdistributor plates could be used, with each distributor plate receivingliquid from one liquid conduit 410 and distributing liquid to a portionof the solids inlet conduits 408. It is further noted that the annularsolids volume could be segmented using one or more internal walls. Thiscould allow a first group of solids inlet conduits 408 to providecatalyst particles to a first portion of the annular solids volume,while a second group of solids inlet conduits provides catalystparticles to a second portion of the annular solids volume. Moregenerally, any convenient number of portions could be defined in theannular solids volume.

Example of Reaction Conditions—Conversion of Oxygenates and/or Olefinsto Naphtha and Distillate

An example of the type of reaction that can be performed using the feeddistribution apparatus, feed separation method, moving bed reactors, andmoving bed reactor configurations described herein is conversion ofoxygenates to olefins, optionally with further oligomerization of theolefins to naphtha and/or distillate boiling range products. Examples ofsuitable oxygenates include feeds containing methanol, dimethyl ether,C₁-C₄ alcohols, ethers with C₁-C₄ alkyl chains, including bothasymmetric ethers containing C₁-C₄ alkyl chains (such as methyl ethylether, propyl butyl ether, or methyl propyl ether) and symmetric ethers(such as diethyl ether, dipropyl ether, or dibutyl ether), orcombinations thereof. It is noted that oxygenates containing at leastone C₁-C₄ alkyl group are intended to explicitly identify oxygenateshaving alkyl groups containing 4 carbons or less. Preferably theoxygenate feed can include at least 30 wt % of one or more suitableoxygenates, or at least 50 wt %, or at least 75 wt %, or at least 90 wt%, or at least 95 wt %. Additionally or alternately, the oxygenate feedcan include at least 50 wt % methanol, such as at least 75 wt %methanol, or at least 90 wt % methanol, or at least 95 wt % methanol. Inparticular, the oxygenate feed can include 30 wt % to 100 wt % ofoxygenate (or methanol), or 50 wt % to 95 wt %, or 75 wt % to 100 wt %,or 75 wt % to 95 wt %. The oxygenate feed can be derived from anyconvenient source. For example, the oxygenate feed can be formed byreforming of hydrocarbons in a natural gas feed to form synthesis gas(H₂, CO, CO₂), and then using the synthesis gas to form methanol (orother alcohols). As another example, a suitable oxygenate feed caninclude methanol, dimethyl ether, or a combination thereof as theoxygenate.

In addition to oxygenates, in some aspects the feed can also includeolefins. In this discussion, the olefins included as part of the feedcan correspond to aliphatic olefins that contain 6 carbons or less, sothat the olefins are suitable for formation of naphtha boiling rangecompounds. The olefin portion of the feed can be mixed with theoxygenates prior to entering a reactor for performing oxygenateconversion, or a plurality of streams containing oxygenates and/orolefins can be mixed within a conversion reactor. The feed can include 5wt % to 40 wt % of olefins (i.e., olefins containing 6 carbons or less),or 5 wt % to 30 wt %, or 10 wt % to 40 wt %, or 10 wt % to 30 wt %. Whenthe conversion is operated under low hydrogen transfer conditions with acatalyst that is selective for formation of paraffins and olefins, theaddition of olefins can allow for further production of paraffins andolefins. In aspects where olefins are included in the feed, the molarratio of oxygenates to olefins can be 20 or less, or 10 or less, or 6.0or less, or 4.0 or less, such as down to a molar ratio of 1.0. Forexample, the molar ratio of oxygenates to olefins can be between 1.0 and20, or between 1.0 and 10, or between 1.0 and 6.0, or between 4.0 and20, or between 6.0 and 20. It is noted that the weight percent ofolefins in the feed can be dependent on the nature of the olefins. Forexample, if a C₅ olefin is used as the olefin with a methanol-containingfeed, the wt % of olefin required to achieve a desired molar ratio ofolefin to oxygenate will be relatively high due to the much largermolecular weight of a C₅ alkene.

In some aspects, the olefins can correspond to olefins generated duringthe oxygenate conversion process. In such aspects, a portion of theeffluent from the conversion process can be recycled to provide olefinsfor the feed. In other aspects, the olefins can be derived from anyother convenient source. The olefin feed can optionally includecompounds that act as inerts or act as a diluent in the conversionprocess. For example, a stream of “waste” olefins having an olefincontent of 5 vol % to 20 vol % can be suitable as a source of olefins,so long as the other components of the “waste” olefins stream arecompatible with the conversion process. For example, the othercomponents of the olefin stream can correspond to inert gases such asN₂, carbon oxides, paraffins, and/or other gases that have lowreactivity under the conversion conditions. Water can also be present,although it can be preferable for water to correspond to 20 vol % orless of the total feed, or 10 vol % or less.

In addition to oxygenates and olefins, a feed can also include diluents,such as water (in the form of steam), nitrogen or other inert gases,and/or paraffins or other non-reactive hydrocarbons. In some aspects,the source of olefins can correspond to a low purity source of olefins,so that the source of olefins corresponds to 20 wt % or less of olefins.In some aspects, the portion of the feed corresponding to componentsdifferent from oxygenates and olefins can correspond to 1 wt % to 60 wt% of the feed, or 1 wt % to 25 wt %, or 10 wt % to 30 wt %, or 20 wt %to 60 wt %. Optionally, the feed can substantially correspond tooxygenates and olefins, so that the content of components different fromoxygenates and olefins is 1 wt % or less (such as down to 0 wt %).

It is noted that the above feed description can correspond to the inputstream for a group of moving bed reactors that are operated in series.In such aspects, the above feed can be passed into a first moving bedreactor, which partially converts oxygenates and/or partiallyoligomerizes olefins. The effluent from the first moving bed reactor canthen be separated into solid particles, gas phase unreacted feed andintermediate products, and liquid products (if any) that have formed.The solid particles, liquid, and gas can then be introduced into asecond moving bed reactor for further reaction. This can be repeateduntil a sufficient number of moving bed reactor stages have been used toachieve desired products. This can correspond to, for example, asufficient number of stages to achieve complete oxygenate conversion, asufficient number of stages so that oligomerization results in a desiredweight percentage (relative to the feed) of distillate boiling rangeproducts, or another convenient reaction end point.

In aspects where a single series of reactors is used to upgradeoxygenates to distillate products, the net yield of C₅₊ hydrocarbons inthe upgrading effluent can be 10 wt % to 90 wt %, or 20 wt % to 80 wt %,or 40 wt % to 90 wt %, or 40 wt % to 80 wt % on a dry basis. Theupgrading effluent can correspond to the effluent from the final movingbed stage of a series of moving bed reactors. The net yield refers tothe yield of C₅₊ hydrocarbons in the upgrading effluent minus the amount(if any) of C₅₊ alkenes in the feed. For example, when pentene is usedas an olefin in the feed, the weight of pentene in the feed issubtracted from the weight of C₅₊ hydrocarbons in the upgrading effluentwhen determining net yield. It is noted that the net yield is expressedon a dry basis due to the high variability in the amount of water thatmay be produced, depending on the oxygenate used as the feed. Forexample, if a pre-conversion stage is used to convert methanol to waterso that dimethyl ether is used as a feed introduced into the moving bedreactors, the weight of water in the conversion effluent can be reducedby roughly 50 wt %. In various aspects, the yield of paraffins plusolefins relative to the C₅₊ portion of the hydrocarbon product can be 20wt % to 90 wt %, or 40 wt % to 90 wt %, or 40 wt % to 80 wt %.Additionally or alternately, the yield of distillate compounds (330°F.+, ˜165° C.+) relative to the C₅₊ portion of the hydrocarbon productcan be 20 wt % to 60 wt %, or 30 wt % to 60 wt %, or 30 wt % to 50 wt %.Further additionally or alternately, less than 10 wt % of the totalhydrocarbon product can correspond to C₁ paraffins (methane).

In other aspects where separate series of reactors are used foroxygenate conversion and oligomerization, the net yield of C₅₊hydrocarbons in the conversion effluent can be 10 wt % to 30 wt % on adry basis. The conversion effluent can correspond to the effluent fromthe final moving bed stage of a first series of moving bed reactors. Insuch aspects, the yield of paraffins plus olefins relative to the C₅₊portion of the hydrocarbon product in the conversion effluent (on a drybasis) can be 20 wt % to 90 wt %, or 40 wt % to 90 wt %, or 40 wt % to80 wt %. Additionally or alternately, the yield of olefins relative tothe total hydrocarbon product in the conversion effluent (on a drybasis) can be 20 wt % to 80 wt %, or 40 wt % to 80 wt %, or 40 wt % to70 wt %. Further additionally or alternately, less than 10 wt % of thetotal hydrocarbon product in the conversion effluent can correspond toC₁ paraffins (methane).

At least a portion of the conversion effluent can then be passed intothe second series of reactors for oligomerization. The oligomerizationeffluent can include a naphtha boiling range portion, a distillate fuelboiling range portion, and a light ends portion. Optionally butpreferably, the oligomerization effluent can include 20 wt % or more ofcompounds boiling above the naphtha boiling range (204° C.+), or 30 wt %or more, on a dry basis.

Suitable and/or effective conditions for performing an oxygenateconversion reaction and/or combined conversion plus oligomerization(i.e. upgrading) can include average reaction temperatures of 230° C. to450° C., 230° C. to 290° C., or 250° C. to 300° C., or 230° C. to 280°C., or 270° C. to 300° C.; total pressures between 1 psig (˜7 kPag) to400 psig (˜2700 kPag), or 10 psig (˜70 kPag) to 150 psig (˜1050 kPag),or 10 psig (˜70 kPag) to 100 psig (˜700 kPag), and an oxygenate spacevelocity between 0.1 hr⁻¹ to 10 hr⁻¹ based on weight of oxygenaterelative to weight of catalyst (WHSV), or 0.1 hr⁻¹ to 5.0 hr⁻¹, or 1.0hr⁻¹ to 5.0 hr⁻¹. In this discussion, average reaction temperature isdefined as the average of the temperature at the reactor inlet and thetemperature at the reactor outlet for the reactor where the conversionreaction is performed. In some aspects, the reaction conditions betweenreactors within a series of reactors can be substantially the same. Insome aspects, the reaction conditions can vary between reactors within aseries of reactors, depending on the reaction taking place in eachmoving bed reactor. For example, in a series of reactors containing fourmoving beds, the first three moving bed reactors can operate at highertemperature to facilitate oxygenate conversion. The final moving bedreactor can operate at a lower temperature to facilitate performingadditional oligomerization on the olefins produced in the first threemoving bed reactors.

In aspects where a separate series of moving bed reactors is used foroligomerization, suitable and/or effective conditions for performing anoligomerization reaction can include average reaction temperatures of180° C. to 250° C., or 200° C. to 250° C.; total pressures between 50psig (˜340 kPag) to 800 psig (˜5,500 kPag), and an oxygenate spacevelocity between 0.1 hr⁻¹ to 10 hr⁻¹ based on weight of oxygenaterelative to weight of catalyst (WHSV), or 0.1 hr⁻¹ to 5.0 hr⁻¹, or 1.0hr⁻¹ to 5.0 hr⁻¹.

In various aspects, a transition metal-enhanced zeolite catalystcomposition can be used for conversion of oxygenate feeds to naphthaboiling range fractions and olefins. In this discussion and the claimsbelow, a zeolite is defined to refer to a crystalline material having aporous framework structure built from tetrahedra atoms connected bybridging oxygen atoms. Examples of known zeolite frameworks are given inthe “Atlas of Zeolite Frameworks” published on behalf of the StructureCommission of the International Zeolite Association”, 6^(th) revisededition, Ch. Baerlocher, L. B. McCusker, D. H. Olson, eds., Elsevier,New York (2007) and the corresponding web site,http://www.iza-structure.org/databases/. Under this definition, azeolite can refer to aluminosilicates having a zeolitic framework typeas well as crystalline structures containing oxides of heteroatomsdifferent from silicon and aluminum. Such heteroatoms can include anyheteroatom generally known to be suitable for inclusion in a zeoliticframework, such as gallium, boron, germanium, phosphorus, zinc, and/orother transition metals that can substitute for silicon and/or aluminumin a zeolitic framework.

A suitable zeolite can include a 1-dimensional or 2-dimensional10-member ring pore channel network. In some aspects, additionalbenefits can be achieved if the zeolite also has 12-member ring pocketsat the surface, such as MWW framework (e.g., MCM-49, MCM-22). Suchpockets represent active sites having a 12-member ring shape, but do notprovide access to a pore network. Examples of MWW framework zeolitesinclude MCM-22, MCM-36, MCM-49, MCM-56, EMM-10, EMM-12, EMM-13, andITQ-2. In some aspects, zeolites with a 1-dimensional or 2-dimensional12-member ring pore channel network can also be suitable, such as MORframework zeolites. Examples of suitable zeolites having a 1-dimensional10-member ring pore channel network include zeolites having a MRE (e.g,ZSM-48), MTW, TON (e.g., ZSM-22), MTT (e.g., ZSM-23), and/or MFSframework. In some aspects, ZSM-48, ZSM-22, MCM-22, MCM-49, or acombination thereof can correspond to preferred zeolites.

Generally, a zeolite having desired activity for methanol conversion canhave a silicon to aluminum molar ratio of 5 to 200, or 15 to 100, or 20to 80, or 20 to 40. For example, the silicon to aluminum ratio can be atleast 10, or at least 20, or at least 30, or at least 40, or at least50, or at least 60. Additionally or alternately, the silicon to aluminumratio can be 300 or less, or 200 or less, or 100 or less, or 80 or less,or 60 or less, or 50 or less.

It is noted that the molar ratio described herein is a ratio of siliconto aluminum. If a corresponding ratio of silica to alumina weredescribed, the corresponding ratio of silica (SiO₂) to alumina (Al₂O₃)would be twice as large, due to the presence of two aluminum atoms ineach alumina stoichiometric unit. Thus, a silicon to aluminum ratio of10 corresponds to a silica to alumina ratio of 20.

In some aspects, a zeolite in a catalyst can be present at least partlyin the hydrogen form. Depending on the conditions used to synthesize thezeolite, this may correspond to converting the zeolite from, forexample, the sodium form. This can readily be achieved, for example, byion exchange to convert the zeolite to the ammonium form followed bycalcination in air or an inert atmosphere at a temperature of 400° C. to700° C. to convert the ammonium form to the active hydrogen form.

Additionally or alternately, a zeolitic catalyst can include and/or beenhanced by a transition metal. The transition metal can be anyconvenient transition metal selected from Groups 6-15 of the IUPACperiodic table. The transition metal can be incorporated into thezeolite/catalyst by any convenient method, such as by impregnation, byion exchange, by mulling prior to extrusion, and/or by any otherconvenient method. Optionally, the transition metal incorporated into azeolite/catalyst can correspond to two or more metals. Afterimpregnation or ion exchange, the transition metal-enhanced catalyst canbe treated in air or an inert atmosphere at a temperature of 400° C. to700° C. The amount of transition metal can be expressed as a weightpercentage of metal relative to the total weight of the catalyst(including any zeolite and any binder). A catalyst can include 0.05 wt %to 20 wt % of one or more transition metals, or 0.1 wt % to 10 wt %, or0.1 wt % to 5 wt %, or 0.1 wt % to 2.0 wt %. For example, the amount oftransition metal can be at least 0.1 wt % of transition metal, or atleast 0.25 wt % of transition metal, or at least 0.5 wt %, or at least0.75 wt %, or at least 1.0 wt %. Additionally or alternately, the amountof transition metal can be 20 wt % or less, or 10 wt % or less, or 5 wt% or less, or 2.0 wt % or less, or 1.5 wt % or less, or 1.2 wt % orless, or 1.1 wt % or less, or 1.0 wt % or less.

A catalyst composition can employ a zeolite in its original crystallineform or after formulation into catalyst particles, such as by extrusion.A process for producing zeolite extrudates in the absence of a binder isdisclosed in, for example, U.S. Pat. No. 4,582,815, the entire contentsof which are incorporated herein by reference. Preferably, thetransition metal can be incorporated after formulation of the zeolite(such as by extrusion) to form catalyst particles without an addedbinder. Optionally, such an “unbound” catalyst can be steamed afterextrusion. The terms “unbound” is intended to mean that the presentcatalyst composition is free of any of the inorganic oxide binders, suchas alumina or silica, frequently combined with zeolite catalysts toenhance their physical properties.

The catalyst compositions described herein can further be characterizedbased on activity for hexane cracking, or Alpha value. Alpha value is ameasure of the acid activity of a zeolite catalyst as compared with astandard silica-alumina catalyst. The alpha test is described in U.S.Pat. No. 3,354,078; in the Journal of Catalysis, Vol. 4, p. 527 (1965);Vol. 6, p. 278 (1966); and Vol. 61, p. 395 (1980), each incorporatedherein by reference as to that description. The experimental conditionsof the test used herein include a constant temperature of 538° C. and avariable flow rate as described in detail in the Journal of Catalysis,Vol. 61, p. 395. Higher alpha values correspond with a more activecracking catalyst. For an oxygenate conversion catalyst, Alpha value canbe 15 to 150, or 15 to 100, or 15 to 50. Lower Alpha values can bebeneficial, as increased acidity can tend to increase hydrogen transfer.In other aspects, such as when the conversion is performed attemperatures of 275° C. or less, or 250° C. or less, catalysts with anAlpha value of 15 to 1000 can be suitable. This is due to thesuppression of hydrogen transfer at lower temperatures.

As an alternative to forming catalysts without a separate binder,zeolite crystals can be combined with a binder to form bound catalysts.Suitable binders for zeolite-based catalysts can include variousinorganic oxides, such as silica, alumina, zirconia, titania,silica-alumina, cerium oxide, magnesium oxide, yttrium oxide, orcombinations thereof. For catalysts including a binder, the catalyst cancomprise at least 10 wt % zeolite, or at least 30 wt %, or at least 50wt %, such as up to 90 wt % or more. Generally, a binder can be presentin an amount between 1 wt % and 90 wt %, for example between 5 wt % and40 wt % of a catalyst composition. In some aspects, the catalyst caninclude at least 5 wt % binder, such as at least 10 wt %, or at least 20wt %. Additionally or alternately, the catalyst can include 90 wt % orless of binder, such as 50 wt % or less, or 40 wt % or less, or 35 wt %or less. Combining the zeolite and the binder can generally be achieved,for example, by mulling an aqueous mixture of the zeolite and binder andthen extruding the mixture into catalyst pellets. A process forproducing zeolite extrudates using a silica binder is disclosed in, forexample, U.S. Pat. No. 4,582,815. Optionally, a bound catalyst can besteamed after extrusion.

ADDITIONAL EMBODIMENTS Embodiment 1

A method for upgrading a feed to form distillate boiling rangecompounds, comprising: passing a catalyst flow comprising a catalystinto a first moving bed reactor of a plurality of serially connectedmoving bed reactors, the first moving bed reactor comprising a firstreactor feed inlet and a first reactor effluent outlet, the catalystcomprising at least one of oxygenate conversion catalyst and olefinoligomerization catalyst; exposing a feed to the catalyst flow in thefirst moving bed reactor under first reaction conditions comprising atleast one of oxygenate conversion conditions and oligomerizationconditions, to form a first partially reacted effluent, a temperaturedifferential between the first reactor feed inlet and the first reactoreffluent outlet being 85° C. or less; stripping the catalyst flow with afirst stripping fluid to separate at least a portion of the firstpartially reacted effluent from the catalyst flow, the at least aportion of the first partially reacted effluent optionally exiting thefirst moving bed reactor through the first reactor effluent outlet, theat least a portion of the first partially reacted effluent comprising aliquid phase effluent portion and a vapor phase effluent portion;passing the stripped catalyst flow into a second moving bed reactor ofthe plurality of serially connected moving bed reactors; passing theliquid phase effluent portion into the second moving bed reactor as asubstantially axial flow; and exposing the vapor phase effluent portionto the stripped catalyst flow in the presence of the liquid effluentportion in the second moving bed reactor under second reactionconditions comprising at least one of oxygenate conversion conditionsand oligomerization conditions, to form a second effluent comprisingdistillate boiling range compounds.

Embodiment 2

The method of Embodiment 1, wherein the feed comprises at least aportion of an effluent from a third moving bed reactor of the pluralityof serially connected moving bed reactors.

Embodiment 3

The method of any of the above embodiments, wherein the catalystcomprises olefin oligomerization catalyst and wherein the first reactionconditions comprise oligomerization conditions, the method furthercomprising: converting an oxygenate feedstock in one or more reactors toform a conversion effluent comprising a conversion total hydrocarbonproduct, the conversion total hydrocarbon product comprising 20 wt % ormore olefins, wherein the feed comprises at least a portion of theconversion total hydrocarbon product.

Embodiment 4

The method of Embodiment 3, wherein converting the oxygenate feedstockcomprises: passing a second catalyst flow comprising oxygenateconversion catalyst into a fourth moving bed reactor of a secondplurality of serially connected moving bed reactors, the fourth movingbed reactor comprising a fourth reactor feed inlet and a fourth reactoreffluent outlet; exposing the oxygenate feedstock to the second catalystflow in the fourth moving bed reactor under oxygenate upgradingconditions to form a partially converted effluent, a temperaturedifferential between the fourth reactor feed inlet and the fourthreactor effluent outlet being 80° C. or less; stripping the secondcatalyst flow with a second stripping fluid to separate at least aportion of the partially converted effluent from the second catalystflow, the at least a portion of the partially converted effluentoptionally exiting the fourth moving bed reactor through the fourthreactor effluent outlet, the at least a portion of the partiallyconverted effluent comprising a vapor phase converted effluent portion;passing the second stripped catalyst flow into a fifth moving bedreactor of the second plurality of serially connected moving bedreactors; and exposing the vapor phase converted effluent portion to thesecond stripped catalyst flow in the fifth moving bed reactor undersecond oxygenate upgrading conditions to form the conversion effluent.

Embodiment 5

The method of Embodiment 4, wherein the at least a portion of thepartially converted effluent further comprises a liquid phase convertedeffluent portion, the method further comprising: passing the liquidphase converted effluent portion into the fifth moving bed reactor as asubstantially axial flow, wherein the vapor phase converted effluentportion is exposed to the second stripped catalyst flow in the fifthmoving bed reactor in the presence of the liquid converted effluentportion.

Embodiment 6

The method of Embodiment 4 or 5, wherein the olefin oligomerizationcatalyst and the oxygenate conversion catalyst are regenerated in acommon regenerator, or wherein a rate of the catalyst flow into thefirst moving bed reactor is different from a rate of the catalyst flowinto the fourth moving bed reactor, or a combination thereof.

Embodiment 7

The method of any of the above embodiments, further comprising:separating spent catalyst from the second effluent; and passing thespent catalyst into a regenerator to form regenerated catalyst, whereinat least a portion of the catalyst flow comprises regenerated catalyst.

Embodiment 8

The method of any of the above embodiments, wherein the strippedcatalyst flow into the second moving bed reactor forms a catalyst bedhaving a top surface comprising one or more cones at an angle of reposeof the catalyst in the stripped catalyst flow, and wherein passing theliquid phase effluent portion into the second moving bed reactorcomprises contacting at least a portion of the liquid phase effluentportion with the one or more cones.

Embodiment 9

The method of any of the above embodiments, wherein the vapor flowwithin the second moving bed reactor comprises an axial vapor flow, orwherein the vapor flow within the second moving bed reactor comprises aradial vapor flow.

Embodiment 10

The method of any of the above embodiments, wherein the oxygenatecatalyst comprises a zeotype catalyst, or wherein oligomerizationcatalyst comprises a zeotype catalyst, or a combination thereof, thezeotype catalyst optionally comprising ZSM-48.

Embodiment 11

A system for upgrading oxygenates, comprising: a first plurality ofserially connected moving bed reactors, each moving bed reactor in thefirst plurality of serially connected moving bed reactors comprising acatalyst inlet, a feed inlet, a stripping fluid inlet, a catalystoutlet, and a vapor phase effluent outlet; a separation stage comprisinga separation stage inlet and one or more separation stage outlets, theseparation stage inlet being in fluid communication with at least atleast one vapor phase effluent outlet; a second plurality of seriallyconnected second moving bed reactors, each moving bed reactor in thesecond plurality of serially connected moving bed reactors comprising asecond catalyst inlet, a second feed inlet, a second stripping fluidinlet, a second catalyst outlet, a second liquid phase effluent outlet,and a second vapor phase effluent outlet, at least one second feed inletbeing in fluid communication with at least one of the one or moreseparation stage outlets; a second separation stage comprising a secondseparation stage inlet and one or more second separation stage outlets,the separation stage inlet being in fluid communication with at leastone second liquid phase effluent outlet and at least one second vaporphase effluent outlet; and a regenerator comprising a first regeneratorinlet, a first regenerator outlet, a second regenerator inlet, a secondregenerator outlet, an oxygen inlet, and a flue gas outlet, the firstregenerator inlet being in solids flow communication with at least onecatalyst outlet, the second regenerator inlet being in solids flowcommunication with at least one second catalyst outlet, the firstregenerator outlet being in solids flow communication with at least onecatalyst inlet, the second regenerator outlet being in solids flowcommunication with at least one second catalyst inlet.

Embodiment 12

The system of Embodiment 11, wherein each moving bed reactor in thefirst plurality of serially connected moving bed reactors furthercomprises a liquid phase effluent outlet, the separation stage being influid communication with at least one liquid phase effluent outlet, andwherein the second feed inlet comprises a gas feed inlet and a liquidfeed inlet.

Embodiment 13

The system of Embodiment 11 or 12, a) wherein one or more moving bedreactors of the first plurality of serially connected moving bedreactors further comprise first solids volumes, the first solids volumescomprising moving beds of oxygenate conversion catalyst having topsurfaces comprising a cone at an angle of repose of the oxygenateconversion catalyst; b) wherein one or more moving bed reactors of thesecond plurality of serially connected moving bed reactors furthercomprise second solids volumes, the second solids volumes comprisingmoving beds of olefin oligomerization catalyst having top surfacescomprising a cone at an angle of repose of the olefin oligomerizationcatalyst; or c) a combination of a) and b).

Embodiment 14

The system of any of Embodiments 11-13, wherein the second plurality ofserially connected moving bed reactors comprises a vertical stack, atleast one second catalyst outlet being in solids flow communication withat least one second catalyst inlet without passing through a catalystflow controller.

Embodiment 15

An effluent comprising distillate boiling range compounds made accordingto the method of any of Embodiments 1-10 or made using the system of anyof Embodiments 11-14.

Additional Embodiment A

The method of any of the above Embodiments 1-10, wherein the at least aportion of the first partially reacted effluent comprises 95 mol % ormore of the hydrocarbonaceous compounds in the first partially reactedeffluent.

Additional Embodiment B

The method of any of Embodiments 1-10, wherein the plurality of seriallyconnected moving bed reactors comprises a vertical stack, and whereinthe stripped catalyst flow is passed into the second reactor withoutpassing through a catalyst flow controller.

Additional Embodiment C

The system of any of Embodiments 11-14, wherein the one or more movingbed reactors of the second plurality of serially connected moving bedreactors comprise one or more of the following elements, or two or moreof the following elements, or a plurality of the following elements, orup to substantially all of the following elements: a reactor comprisingan annular outer volume, an annular solids volume inside the annularouter volume, a first central conduit inside of the annular solidsvolume, and an inner central conduit inside the first central conduit,the annular solids volume comprising a plurality of perforationsproviding vapor communication between the annular outer volume and thefirst central conduit, the perforations preventing flow of solidparticles into the annular outer volume and into the first centralconduit; a central gas opening in fluid communication with the outerannular volume; a plurality of solids inlet conduits in solids flowcommunication with the annular solids volume; one or more distributorplates comprising distributor plate concave volumes, a plurality of thedistributor plate concave volumes being arranged around each of thesolids inlet conduits, each distributor plate concave volume comprisingone or more orifices providing fluid communication between eachdistributor plate concave volume and the annular solids volume; aplurality of liquid inlet conduits in fluid communication with thedistributor plate concave volumes; a gas exit conduit in fluidcommunication with the inner central volume; a liquid exit conduit influid communication with the outer central volume; and a solids exitvolume in solids flow communication with a bottom of the solids annularvolume, the solids exit volume further comprising a stripping gas inletand a stripping gas outlet, the stripping gas outlet providing fluidcommunication with the outer central volume.

Additional Embodiment D

The system of any of Embodiments 11-14 or Additional Embodiment C,wherein the one or more moving bed reactors of the second plurality ofserially connected moving bed reactors comprise one or more of thefollowing elements, or two or more of the following elements, or all ofthe following elements: one or more solids inlet conduits in solids flowcommunication with a solids volume in a moving bed reactor, the one ormore solids inlet conduits having a smaller characteristic width than awidth of the solids volume at an interface between each solids inletconduit and the solids volume; a distributor plate comprising aplurality of concave volumes around each of the one or more solids inletconduits, and a plurality of exit surfaces separating the plurality ofconcave volumes from the solids volume within the moving bed reactor,each concave volume comprising one or more orifices providing fluidcommunication, through an exit surface, between the concave volume andthe solids volume; and a plurality of fluid inlet conduits in fluidcommunication with the plurality of concave volumes, wherein theplurality of exit surfaces are oriented at an angle of 15° to 45°relative to a plane defined by at least one interface between the one ormore solids inlet conduits and the solids volume.

While the present invention has been described and illustrated byreference to particular embodiments, those of ordinary skill in the artwill appreciate that the invention lends itself to variations notnecessarily illustrated herein. For this reason, then, reference shouldbe made solely to the appended claims for purposes of determining thetrue scope of the present invention.

1. A method for upgrading a feed to form distillate boiling rangecompounds, comprising: passing a catalyst flow comprising a catalystinto a first moving bed reactor of a plurality of serially connectedmoving bed reactors, the first moving bed reactor comprising a firstreactor feed inlet and a first reactor effluent outlet, the catalystcomprising at least one of oxygenate conversion catalyst and olefinoligomerization catalyst; exposing a feed to the catalyst flow in thefirst moving bed reactor under first reaction conditions comprising atleast one of oxygenate conversion conditions and oligomerizationconditions, to form a first partially reacted effluent, a temperaturedifferential between the first reactor feed inlet and the first reactoreffluent outlet being 85° C. or less; stripping the catalyst flow with afirst stripping fluid to separate at least a portion of the firstpartially reacted effluent from the catalyst flow, the at least aportion of the first partially reacted effluent comprising a liquidphase effluent portion and a vapor phase effluent portion; passing thestripped catalyst flow into a second moving bed reactor of the pluralityof serially connected moving bed reactors; passing the liquid phaseeffluent portion into the second moving bed reactor as a substantiallyaxial flow; and exposing the vapor phase effluent portion to thestripped catalyst flow in the presence of the liquid effluent portion inthe second moving bed reactor under second reaction conditionscomprising at least one of oxygenate conversion conditions andoligomerization conditions, to form a second effluent comprisingdistillate boiling range compounds.
 2. The method of claim 1, whereinthe feed comprises at least a portion of an effluent from a third movingbed reactor of the plurality of serially connected moving bed reactors.3. The method of claim 1, wherein the at least a portion of the firstpartially reacted effluent exits the first moving bed reactor throughthe first reactor effluent outlet.
 4. The method of claim 1, wherein thecatalyst comprises olefin oligomerization catalyst and wherein the firstreaction conditions comprise oligomerization conditions, the methodfurther comprising: converting an oxygenate feedstock in one or morereactors to form a conversion effluent comprising a conversion totalhydrocarbon product, the conversion total hydrocarbon product comprising20 wt % or more olefins, wherein the feed comprises at least a portionof the conversion total hydrocarbon product.
 5. The method of claim 4,wherein converting the oxygenate feedstock comprises: passing a secondcatalyst flow comprising oxygenate conversion catalyst into a fourthmoving bed reactor of a second plurality of serially connected movingbed reactors, the fourth moving bed reactor comprising a fourth reactorfeed inlet and a fourth reactor effluent outlet; exposing the oxygenatefeedstock to the second catalyst flow in the fourth moving bed reactorunder oxygenate upgrading conditions to form a partially convertedeffluent, a temperature differential between the fourth reactor feedinlet and the fourth reactor effluent outlet being 80° C. or less;stripping the second catalyst flow with a second stripping fluid toseparate at least a portion of the partially converted effluent from thesecond catalyst flow, the at least a portion of the partially convertedeffluent comprising a vapor phase converted effluent portion; passingthe second stripped catalyst flow into a fifth moving bed reactor of thesecond plurality of serially connected moving bed reactors; and exposingthe vapor phase converted effluent portion to the second strippedcatalyst flow in the fifth moving bed reactor under second oxygenateupgrading conditions to form the conversion effluent.
 6. The method ofclaim 5, wherein the at least a portion of the partially convertedeffluent exits the fourth moving bed reactor through the fourth reactoreffluent outlet.
 7. The method of claim 5, wherein the at least aportion of the partially converted effluent further comprises a liquidphase converted effluent portion, the method further comprising: passingthe liquid phase converted effluent portion into the fifth moving bedreactor as a substantially axial flow, wherein the vapor phase convertedeffluent portion is exposed to the second stripped catalyst flow in thefifth moving bed reactor in the presence of the liquid phase convertedeffluent portion.
 8. The method of claim 5, wherein the olefinoligomerization catalyst and the oxygenate conversion catalyst areregenerated in a common regenerator.
 9. The method of claim 5, wherein arate of the catalyst flow into the first moving bed reactor is differentfrom a rate of the catalyst flow into the fourth moving bed reactor. 10.The method of claim 1, further comprising: separating spent catalystfrom the second effluent; and passing the spent catalyst into aregenerator to form regenerated catalyst, wherein at least a portion ofthe catalyst flow comprises regenerated catalyst.
 11. The method ofclaim 1, wherein the at least a portion of the first partially reactedeffluent comprises 95 mol % or more of the hydrocarbonaceous compoundsin the first partially reacted effluent.
 12. The method of claim 1,wherein the stripped catalyst flow into the second moving bed reactorforms a catalyst bed having a top surface comprising one or more conesat an angle of repose of the catalyst in the stripped catalyst flow, andwherein passing the liquid phase effluent portion into the second movingbed reactor comprises contacting at least a portion of the liquid phaseeffluent portion with the one or more cones.
 13. The method of claim 1,wherein the vapor flow within the second moving bed reactor comprises anaxial vapor flow, or wherein the vapor flow within the second moving bedreactor comprises a radial vapor flow.
 14. The method of claim 1,wherein the oxygenate catalyst comprises a zeotype catalyst, or whereinoligomerization catalyst comprises a zeotype catalyst, or a combinationthereof.
 15. The method of claim 14, wherein the zeotype catalystcomprises ZSM-48.
 16. The method of claim 1, wherein the plurality ofserially connected moving bed reactors comprises a vertical stack, andwherein the stripped catalyst flow is passed into the second reactorwithout passing through a catalyst flow controller.
 17. A system forupgrading oxygenates, comprising: a first plurality of seriallyconnected moving bed reactors, each moving bed reactor in the firstplurality of serially connected moving bed reactors comprising acatalyst inlet, a feed inlet, a stripping fluid inlet, a catalystoutlet, and a vapor phase effluent outlet; a separation stage comprisinga separation stage inlet and one or more separation stage outlets, theseparation stage inlet being in fluid communication with at least atleast one vapor phase effluent outlet; a second plurality of seriallyconnected second moving bed reactors, each moving bed reactor in thesecond plurality of serially connected moving bed reactors comprising asecond catalyst inlet, a second feed inlet, a second stripping fluidinlet, a second catalyst outlet, a second liquid phase effluent outlet,and a second vapor phase effluent outlet, at least one second feed inletbeing in fluid communication with at least one of the one or moreseparation stage outlets; a second separation stage comprising a secondseparation stage inlet and one or more second separation stage outlets,the separation stage inlet being in fluid communication with at leastone second liquid phase effluent outlet and at least one second vaporphase effluent outlet; and a regenerator comprising a first regeneratorinlet, a first regenerator outlet, a second regenerator inlet, a secondregenerator outlet, an oxygen inlet, and a flue gas outlet, the firstregenerator inlet being in solids flow communication with at least onecatalyst outlet, the second regenerator inlet being in solids flowcommunication with at least one second catalyst outlet, the firstregenerator outlet being in solids flow communication with at least onecatalyst inlet, the second regenerator outlet being in solids flowcommunication with at least one second catalyst inlet.
 18. The system ofclaim 17, wherein each moving bed reactor in the first plurality ofserially connected moving bed reactors further comprises a liquid phaseeffluent outlet, the separation stage being in fluid communication withat least one liquid phase effluent outlet.
 19. The system of claim 17,wherein the second feed inlet comprises a gas feed inlet and a liquidfeed inlet.
 20. The system of claim 17, wherein one or more moving bedreactors of the first plurality of serially connected moving bedreactors further comprise first solids volumes, the first solids volumescomprising moving beds of oxygenate conversion catalyst having topsurfaces comprising a cone at an angle of repose of the oxygenateconversion catalyst.
 21. The system of claim 17, wherein one or moremoving bed reactors of the second plurality of serially connected movingbed reactors further comprise second solids volumes, the second solidsvolumes comprising moving beds of oxygenate conversion catalyst havingtop surfaces comprising a cone at an angle of repose of the oxygenateconversion catalyst.
 22. The system of claim 17, wherein the secondplurality of serially connected moving bed reactors comprises a verticalstack, at least one second catalyst outlet being in solids flowcommunication with at least one second catalyst inlet without passingthrough a catalyst flow controller.
 23. The system of claim 17, whereinthe one or more moving bed reactors of the second plurality of seriallyconnected moving bed reactors comprise: a reactor comprising an annularouter volume, an annular solids volume inside the annular outer volume,a first central conduit inside of the annular solids volume, and aninner central conduit inside the first central conduit, the annularsolids volume comprising a plurality of perforations providing vaporcommunication between the annular outer volume and the first centralconduit, the perforations preventing flow of solid particles into theannular outer volume and into the first central conduit; a central gasopening in fluid communication with the outer annular volume; aplurality of solids inlet conduits in solids flow communication with theannular solids volume; one or more distributor plates comprisingdistributor plate concave volumes, a plurality of the distributor plateconcave volumes being arranged around each of the solids inlet conduits,each distributor plate concave volume comprising one or more orificesproviding fluid communication between each distributor plate concavevolume and the annular solids volume; a plurality of liquid inletconduits in fluid communication with the distributor plate concavevolumes; a gas exit conduit in fluid communication with the innercentral volume; a liquid exit conduit in fluid communication with theouter central volume; and a solids exit volume in solids flowcommunication with a bottom of the solids annular volume, the solidsexit volume further comprising a stripping gas inlet and a stripping gasoutlet, the stripping gas outlet providing fluid communication with theouter central volume.
 24. The system of claim 17, wherein the one ormore moving bed reactors of the second plurality of serially connectedmoving bed reactors comprise: one or more solids inlet conduits insolids flow communication with a solids volume in a moving bed reactor,the one or more solids inlet conduits having a smaller characteristicwidth than a width of the solids volume at an interface between eachsolids inlet conduit and the solids volume; a distributor platecomprising a plurality of concave volumes around each of the one or moresolids inlet conduits, and a plurality of exit surfaces separating theplurality of concave volumes from the solids volume within the movingbed reactor, each concave volume comprising one or more orificesproviding fluid communication, through an exit surface, between theconcave volume and the solids volume; and a plurality of fluid inletconduits in fluid communication with the plurality of concave volumes,wherein the plurality of exit surfaces are oriented at an angle of 15°to 45° relative to a plane defined by at least one interface between theone or more solids inlet conduits and the solids volume.